Process for preparing polyolefin products

ABSTRACT

A novel liquid phase polymerization process for preparing a polyolefin product having preselected properties is disclosed. The process includes the steps of providing a liquid feedstock which contains an olefinic component and a catalyst composition consisting of a stable complex of BF 3  and a complexing agent therefor. The feedstock may comprise any one or more of a number of olefins including branched olefins such as isobutylene, C 3  to C 15  linear alpha olefins and C 4  to C 15  reactive non-alpha olefins. The feedstock and the catalyst composition are introduced into a residual reaction mixture recirculating in a loop reactor reaction zone provided in the tube side of a shell and tube heat exchanger at a recirculation rate sufficient to cause intimate intermixing of the residual reaction mixture, the added feedstock and the added catalyst composition. The heat of the polymerization reaction is removed from the recirculating intimately intermixed reaction admixture at a rate calculated to provide a substantially constant reaction temperature therein while the same is recirculating in said reaction zone. The conditions in the reactor are appropriate for causing olefinic components introduced in said feedstock to undergo polymerization to form the desired polyolefin product in the presence of the catalyst composition. A product stream containing the desired polyolefin product is withdrawn from the reaction zone. The introduction of the feedstock into the reaction zone and the withdrawal of the product stream from the reaction zone are controlled such that the residence time of the olefinic components undergoing polymerization in the reaction zone is appropriate for production of the desired polyolefin product.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a continuation of utility application Ser. No.10/208,234 filed Jul. 30, 2002, now U.S. Pat. No. 6,884,858, which inturn is a continuation of utility application Ser. No. 09/665,084 filedon Sep. 20, 2000, now U.S. Pat. No. 6,525,149, which again in turn is acontinuation-in-part of utility application Ser. No. 09/515,790 filed onFeb. 29, 2000, now U.S. Pat. No. 6,562,913. Priority from each of saidcopending utility applications is claimed herein pursuant to 35 U.S.C. §120. In addition, priority benefits under 35 U.S.C. §119(e) are claimedin this application from provisional application Ser. No. 60/160,357,filed on Oct. 19, 1999. The entireties of the disclosures of said priorapplications are hereby specifically incorporated herein by reference.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to olefin polymerization and to thepreparation of polyolefin products. In particular the present inventionrelates to the preparation of a variety of polyolefin products using aliquid phase polymerization process. In this latter regard, theinvention relates to a novel liquid phase process for the polymerizationof olefins using a modified BF₃ catalyst which is stabilized with acomplexing agent.

2. The Prior Art Background

The polymerization of olefins using Friedel-Crafts type catalysts,including BF₃, is a generally known procedure. The degree ofpolymerization of the products obtained varies according to which of thevarious known polymerization techniques is used. In this latter regard,it is to be understood that the molecular weight of the polymericproduct is directly related to the degree of polymerization and that thedegree of polymerization may be manipulated by manipulating processparameters so as to produce a variety of products having respectivedesired average molecular weights.

Generally speaking, due to the nature and mechanics of the olefinicpolymerization process, a polyolefin product has a single double bondremaining in each molecule at the end of the polymerization process. Theposition of this remaining double bond is often an important feature ofthe product. For example, polyisobutylene (PIB) molecules wherein theremaining double bond is in a terminal (vinylidene) position are knownto be more reactive than PIB molecules wherein the remaining double bondis internal, that is, not in a terminal position. A PIB product whereinat least 50% of the double bonds are in a terminal position may often bereferred to as high vinylidene or highly reactive PIB. The extent towhich a polyolefin product has terminal double bonds may also bemanipulated by manipulation of process parameters.

Current processes for olefin oligomerization often employBF₃/co-catalyst systems wherein the BF₃ is complexed with a co-catalyst.This is done for a variety of reasons that are well known to thoseskilled in the olefin polymerization field. For example, and as isexplained in U.S. Pat. No. 5,408,018, a complexed BF₃ catalyst may beuseful for manipulating and attempting to balance the molecular weight,vinylidene content and polydispersity of PIB. The co-catalyst often ispropanol or a higher alcohol and such co-catalyst systems are usedirrespective of whether the desired product is a poly alpha olefin or apoly internal olefin. However, the use of alcohols having beta hydrogenatoms in such co-catalyst complexes is troublesome because, over time,the BF₃ tends to attack the beta hydrogen atoms. This leads todecomposition of the alcohol whereby the catalyst is renderedineffective. Thus, the co-catalyst complex is unstable and often has avery short shelf life.

To address this problem, many current processes employ a procedurewhereby the co-catalyst complex is prepared in-situ by mixing thealcohol and gaseous BF₃ immediately prior to introduction of theco-catalyst complex into a reactor. In addition, it is not unusual inthe conduct of processes employing such co-catalyst systems to use anexcess of alcohol and to sparge gaseous BF₃ into the reaction mass atseveral downstream points to rebuild catalyst activity. Such methodologyimplies a three-phase reaction and the necessity of using a stirred tankreactor to provide means of dispersing gaseous BF₃ into the reactionmass. These processes use either batch reactors or a set of continuouslystirred tank reactors in series to provide both gas handling capabilityand to satisfy the necessity for a plug flow reactor configuration.

It is also known that alpha olefins, particularly PIB, may bemanufactured in at least two different grades—regular and highvinylidene. Conventionally, these two product grades have been made bydifferent processes, but both often and commonly use a diluted feedstockin which the isobutylene concentration may range from 40-60% by weight.More recently it has been noted that at least the high vinylidene PIBmay be produced using a concentrated feedstock having an isobutylenecontent of 90% by weight or more. Non-reactive hydrocarbons, such asisobutane, n-butane and/or other lower alkanes commonly present inpetroleum fractions, may also be included in the feedstock as diluents.The feedstock often may also contain small quantities of otherunsaturated hydrocarbons such as 1-butene and 2-butene.

Regular grade PIB may range in molecular weight from 500 to 1,000,000 orhigher, and is generally prepared in a batch process at low temperature,sometimes as low as −50 to −70° C. AlCl₃, RAlCl₂ or R₂AlCl are used ascatalysts. The catalyst is not totally removed from the final PIBproduct. Molecular weight may be controlled by temperature since themolecular weight of the product varies inversely with temperature. Thatis to say, higher temperatures give lower molecular weights. Reactiontimes are often in the order of hours. The desired polymeric product hasa single double bond per molecule, and the double bonds are mostlyinternal. Generally speaking, at least about 90% of the double bonds areinternal and less than 10% of the double bonds are in a terminalposition. Even though the formation of terminal double bonds is believedto be kinetically favored, the long reaction times and the fact that thecatalyst is not totally removed, both favor the rearrangement of themolecule so that the more thermodynamically favored internal double bondisomers are formed. Regular PIB may be used as a viscosity modifier,particularly in lube oils, as a thickener, and as a tackifier forplastic films and adhesives. PIB can also be functionalized to produceintermediates for the manufacture of detergents and dispersants forfuels and lube oils.

High vinylidene PIB, a relatively new product in the marketplace, ischaracterized by a large percentage of terminal double bonds, typicallygreater than 70% and preferentially greater than 80%. This provides amore reactive product, compared to regular PIB, and hence this productis also referred to as highly reactive PIB. The terms highly reactive(HR-PIB) and high vinylidene (HV-PIB) are synonymous. The basicprocesses for producing HV-PIB all include a reactor system, employingBF₃ and/or modified BF₃ catalysts, such that the reaction time can beclosely controlled and the catalyst can be immediately neutralized oncethe desired product has been formed. Since formation of the terminaldouble bond is kinetically favored, short reactions times favor highvinylidene levels. The reaction is quenched, usually with an aqueousbase solution, such as, for example, NH₄OH, before significantisomerization to internal double bonds can take place. Molecular weightsare relatively low. HV-PIB having an average molecular weight of about950-1050 is the most common product. Conversions, based on isobutylene,are kept at 75-85%, since attempting to drive the reaction to higherconversions reduces the vinylidene content through isomerization. PriorU.S. Pat. No. 4,152,499 dated May 1, 1979, U.S. Pat. No. 4,605,808 datedAug. 12, 1986, U.S. Pat. No. 5,068,490 dated Nov. 26, 1991, U.S. Pat.No. 5,191,044 dated Mar. 2, 1993, U.S. Pat. No. 5,286,823 dated Jun. 22,1992, U.S. Pat. No. 5,408,018 dated Apr. 18, 1995 and U.S. Pat. No.5,962,604 dated Oct. 5, 1999 are all directed to related subject matter.

U.S. Pat. No. 4,152,499 describes a process for the preparation of PIBsfrom isobutylene under a blanket of gaseous BF₃ acting as apolymerization catalyst. The process results in the production of a PIBwherein 60 to 90% of the double bonds are in a terminal (vinylidene)position.

U.S. Pat. No. 4,605,808 discloses a process for preparing PIB wherein acatalyst consisting of a complex of BF₃ and an alcohol is employed. Itis suggested that the use of such a catalyst complex enables moreeffective control of the reaction parameters. Reaction contact times ofat least 8 minutes are required to obtain a PIB product wherein at leastabout 70% of the double bonds are in a terminal position.

U.S. Pat. No. 5,191,044 discloses a PIB production process requiringcareful pretreatment of a BF₃/alcohol complex to insure that all freeBF₃ is absent from the reactor. The complex must contain a surplus ofthe alcohol complexing agent in order to obtain a product wherein atleast about 70% of the double bonds are in a terminal position. The onlyreaction time exemplified is 10 minutes, and the reaction is carried outat temperatures below 0° C.

In addition to close control of reaction time, the key to obtaining highvinylidene levels seems to be control of catalyst reactivity. This hasbeen done in the past by complexing BF₃ with various oxygenatesincluding sec-butanol and MTBE. One theory is that these complexes areactually less reactive than BF₃ itself, disproportionately slowing theisomerization reaction and thus allowing for greater differentiationbetween the vinylidene forming reaction (polymerization) and theisomerization reaction rates. Mechanisms have also been proposed thatsuggest the BF₃ complexes are non-protonated and thus are not capable ofisomerizing the terminal double bond. This further suggests that water(which can preferentially protonate BF₃) must generally be excluded fromthese reaction systems. In fact, prior publications describingpreparation of PIB using BF₃ complexes teach low water feed (less than20 ppm) is critical to formation of the high vinylidene product.

HV-PIB is increasingly replacing regular grade PIB for the manufactureof intermediates, not only because of higher reactivity, but alsobecause of developing requirements for “chloride free” materials in thefinal product applications. Important PIB derivatives are PIB amines,PIB alkylates and PIB maleic anhydride adducts.

PIB amines can be produced using a variety of procedures involvingdifferent PIB intermediates which provide a reactive site for subsequentamination. These intermediates may include, for example, epoxides,halides, maleic anhydride adducts, and carbonyl derivatives.

Reference to HV-PIB as “highly reactive” is relative to regular gradePIB. HV-PIB is still not, in absolute terms, highly reactive towardformation of some of these intermediates. Other classes of compounds,polyethers for example, can be much more reactive in the formation ofamines and amine intermediates. Amines derived from polyethers are knownas polyether amines (PEAs) and are competitive products to PIB amines.

The use of HV-PIB as an alkylating agent for phenolic compounds, istriggered by the higher reactivity and higher yields achievable withHV-PIB. These very long chain alkyl phenols are good hydrophobes forsurfactants and similar products.

The largest volume PIB derivatives are the PIB-maleic anhydride reactionproducts. HV-PIB is reacted with maleic anhydride through the doublebond giving a product with anhydride functionality. This functionalityprovides reactivity for the formation of amides and other carboxylatederivatives. These products are the basis for most of the lube oildetergents and dispersants manufactured today. As mentioned above,PIB-maleic anhydride products can also be used as intermediates in themanufacture of PIB amine fuel additives.

Other polyolefins which are commercially useful for a variety ofpurposes include conventional PIB wherein the vinylidene content is lessthan 50%, low molecular weight (<350 and perhaps even <250) oligomers ofbranched monomers such as isobutylene, oligomers and higher molecularweight polymers of linear C₃-C₁₅ alpha olefins, and oligomers and highermolecular weight polymers of linear C₄-C₁₅ non-alpha (internal doublebond) olefins. Although these materials are all well known to thoseskilled in the olefin polymerization field, there is always a need fornew developments which improve process efficiency and/or productqualities and reduce operating costs and/or capital expenditures.

SUMMARY OF THE INVENTION

The present invention provides a novel process for the efficient andeconomical production of polyolefin products. Generally speaking, theinvention provides a liquid phase polymerization process for preparing apolyolefin product having preselected properties. In accordance with theprinciples and concepts of the invention, the process includes the stepsof providing a liquid feedstock comprising at least one olefiniccomponent and a catalyst composition comprising a stable complex of BF₃and a complexing agent therefor. The feedstock and the catalystcomposition are introduced into a residual reaction mixture in a loopreactor reaction zone where the residual reaction mixture isrecirculated at a recirculation rate sufficient to cause intimateintermixing of the residual reaction mixture, the added feedstock andthe added catalyst composition to thereby present a recirculating,intimately intermixed reaction admixture in said reaction zone. Therecirculating intimately intermixed reaction admixture is maintained inits intimately intermixed condition while the heat of reaction isremoved therefrom at a rate calculated to provide a substantiallyconstant reaction temperature in the reaction admixture while the sameis recirculating in said reaction zone. The constant reactiontemperature is at a level appropriate for causing olefinic componentsintroduced in said feedstock to undergo polymerization to form thedesired polyolefin product in the presence of the catalyst composition.A product stream comprising the desired polyolefin product is withdrawnfrom the reaction zone. In accordance with the invention, theintroduction of the feedstock into the reaction zone and the withdrawalof the product stream from the reaction zone are controlled such thatthe residence time of the olefinic components undergoing polymerizationin the reaction zone is appropriate for production of the desiredpolyolefin product.

In accordance with one preferred form of the invention, the reactionzone may comprises the tube side of a shell-and-tube heat exchanger. Theheat of the exothermic olefin polymerization reaction may be removedsimultaneously with its generation by circulation of a coolant in theshell side of the exchanger. Preferably, the residence time of theolefinic components undergoing polymerization may be no greater thanabout 3 minutes. Even more preferably, such residence time may be nogreater than about 2 minutes. More preferably still, such residence timemay be no greater than about 1 minute. Ideally, the residence time maybe less than 1 minute.

In accordance with another preferred form of the invention, thecomplexing agent should preferably be such the a stable catalyst complexis formed with BF₃. This is particularly advantageous at the relativelyhigh reaction temperatures needed for oligomerization processes. In thisregard, the complexing agent may advantageously comprise an alcohol,preferably a primary alcohol, and even more preferably a C₁-C₈ primaryalcohol. In a highly preferred form of the invention, the alcohol shouldhave no hydrogen atom on a β carbon. In this highly preferred form ofthe invention, the alcohol may be, for example, methanol or neopentanol.

In accordance with yet another preferred form of the invention, thecomplexing agent may comprise a glycol, preferably glycol wherein eachhydroxyl group of the glycol is in a primary position, and even morepreferably a C₁-C₈ glycol wherein each hydroxyl group of the glycol isin a primary position. In this highly preferred form of the invention,the glycol may be, for example, ethylene glycol.

In conformity with the concepts and principles of another aspect of theinvention, the molar ratio of BF₃ to complexing agent in catalystcomplex may range from approximately 0.5:1 to approximately 5:1.Preferably the molar ratio of BF₃ to complexing agent in said complexmay range from approximately 0.5:1 to approximately 2:1. Even morepreferably, the molar ratio of BF₃ to complexing agent in the complexmay range from approximately 0.5:1 to approximately 1:1. Ideally, themolar ratio of BF₃ to complexing agent in complex may be approximately1:1. Alternatively, the molar ratio of BF₃ to complexing agent in saidcomplex may be approximately 0.75:1.

According to another aspect of the invention, the process may desirablybe conducted such that from about 0.1 to about 10 millimoles of BF₃ areintroduced into the reaction admixture with said catalyst compositionfor each mole of olefinic component introduced into said admixture insaid feedstock. Preferably, from about 0.5 to about 2 millimoles of BF₃may be introduced into the reaction admixture with the catalystcomposition for each mole of olefinic component introduced into theadmixture in said feedstock.

Another important preferred feature of the invention involves thecontinuous recirculation of the reaction admixture at a first volumetricflow rate, and the continuous introduction of the feedstock and thecatalyst composition at a combined second volumetric flow rate.Desirably the ratio of the first volumetric flow rate to the secondvolumetric flow rate may range from about 20:1 to about 50:1. Preferablythe ratio of the first volumetric flow rate to the second volumetricflow rate may range from about 25:1 to about 40:1. Ideally the ratio ofthe first volumetric flow rate to the second volumetric flow rate mayrange from about 28:1 to about 35:1. With regard to this latter aspectof the invention, the ratio of the first volumetric flow rate to thesecond volumetric flow rate may be such that the concentrations ofingredients in the reaction admixture remain essentially constant andsuch that essentially isothermal conditions are established andmaintained in said reaction admixture.

In conformity with the principles and concepts of the invention, thefeedstock and the catalyst composition may be premixed and introducedinto the reaction zone together as a single stream at said secondvolumetric flow rate. Alternatively, the feedstock and the catalystcomposition may be introduced into the reaction zone separately as twostreams, the flow rates of which together add up to said secondvolumetric flow rate.

In further conformity with the principles and concepts of the invention,the reactor configuration, the properties of the reaction mixture, andthe first volumetric flow rate may preferably be such that turbulentflow is maintained in said reaction zone. In this regard, in an idealform of the invention, a Reynolds number of at least about 2000 ismaintained in said reaction zone. In still further conformity with theprinciples and concepts of the invention, the reactor may take the formof the tube side of a shell-and-tube heat exchanger. In this regard, inan ideal form of the invention, a U of at least about 50 Btu/min ft²° F.is maintained in reaction zone.

Preferably, in accordance with the invention, the feed stock maycomprise at least about 30% by weight of said olefinic component.Additionally, the feed stock may include non-reactive hydrocarbondiluents. In this latter regard, the feed stock may comprise at leastabout 30% by weight of said olefinic component with the remainder beingnon-reactive hydrocarbon diluents.

The polymerization process of the invention may be a cationic process.Alternatively the polymerization process of the invention may be acovalent process. An important feature of the invention is that thepolyolefin product of the process of the invention may have a molecularweight of at least about 350 but no more than about 5000. Alternatively,the polyolefin product of the process of the invention may have amolecular weight no greater than about 350 and perhaps no greater thanabout 250.

In accordance with an important aspect of the invention, the olefiniccomponent which is subjected to polymerization may comprise isobutyleneand the polyolefin product may comprise PIB. In further accordance withthis aspect of the invention, the PIB may have a vinylidene content ofat least about 50%. Alternatively, the PIB may have a vinylidene contentno greater than about 50%.

In accordance with yet another important aspect of the invention, theolefinic component may be a branched compound and the product maycomprise a two, three or four member oligomer. The olefinic componentused in the process of the invention may comprise isobutylene and thepolyolefin product may comprise a C₁₂, C₁₆, C₂₀, or C₂₄ PIB oligomer.Alternatively, the olefinic component may comprise either a C₃ to C₁₅linear alpha olefin or a C₄ to C₁₅ reactive non-alpha olefin such as2-butene.

The present invention further provides a novel process for the efficientand economical production of HV-PIB. Generally speaking, the inventionprovides a HV-PIB production process wherein the polymerization reactiontakes place at higher temperatures and at lower reaction times than werethought possible in the past. In particular, the present inventionprovides a liquid phase polymerization process for preparing lowmolecular weight, highly reactive polyisobutylene. Generally speaking,the process may involve cationic polymerization. However, under someconditions the polymerization reaction may be covalent. Particularly thelatter may be true when ether is used as a complexing agent. Inaccordance with this embodiment of the invention, the process includesthe provision of a feedstock comprising isobutylene and a catalystcomposition comprising a complex of BF₃ and a complexing agent. Thefeedstock and the catalyst composition are introduced either separatelyor as a single mixed stream into a residual reaction mixture in areaction zone. The residual reaction mixture, the feedstock and thecatalyst composition are then intimately intermixed to present anintimately intermixed reaction admixture in said reaction zone. Thereaction admixture is maintained in its intimately intermixed conditionand kept at a temperature of at least about 0° C. while the same is insaid reaction zone, whereby the isobutylene in the reaction admixture iscaused to undergo polymerization to form a polyisobutylene product. Aproduct stream comprising a low molecular weight, highly reactivepolyisobutylene is then withdrawn from the reaction zone. Theintroduction of the feedstock into said reaction zone and the withdrawalof the product stream from the reaction zone are controlled such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is no greater than about 4 minutes. In accordance with theinvention, it is possible to conduct the reaction so that the residencetime is no greater than about 3 minutes, no greater than about 2minutes, no greater than about 1 minute, and ideally, even less than 1minute.

In accordance with the concepts and principles of the invention, theprocess may be conducted in a manner such that the polyisobutylene thusproduced has a molecular weight in the range of from about 250 to about5000, in the range of from about 600 to about 4000, in the range of fromabout 700 to about 3000, in the range of from about 800 to about 2000,and ideally in the range of from about 950 to about 1050. In accordancewith the invention, it is possible to so control the process that aparticular molecular weight, such as for example, a molecular weight ofabout 1000, may be achieved.

A major purpose of the invention is to provide a process which may becontrolled sufficiently to insure the production of a polyisobutyleneproduct having a vinylidene content of at least about 70%. Morepreferably the PIB product may have a vinylidene content of at leastabout 80%. Vinylidene contents of at least about 90% may also beachieved through the use of the invention.

The complexing agent used to complex with the BF₃ catalyst may desirablybe an alcohol, and preferably may be a primary alcohol. More preferablythe complexing agent may comprise a C₁-C₈ primary alcohol and ideallymay be methanol.

To achieve the desired results of the invention, the molar ratio of BF₃to complexing agent in the complex may range from approximately 0.5:1 toapproximately 5:1. Preferably the molar ratio of BF₃ to complexing agentin the complex may range from approximately 0.5:1 to approximately 2:1.Even more preferably the molar ratio of BF₃ to complexing agent in thecomplex may range from approximately 0.5:1 to approximately 1:1, andideally, the molar ratio of BF₃ to complexing agent in the complex maybe approximately 1:1.

According to the principles and concepts of the invention, it ispreferred that from about 0.1 to about 10 millimoles of BF₃ may beintroduced into the reaction admixture with the catalyst composition foreach mole of isobutylene introduced into the admixture in the feedstock.Even more preferably, from about 0.5 to about 2 millimoles of BF₃ may beintroduced into the reaction admixture with said catalyst compositionfor each mole of isobutylene introduced into the admixture in thefeedstock.

The invention provides a process whereby the polydispersity of saidpolyisobutylene may be no more than about 2.0, and desirably may be nomore than about 1.65. Ideally, the polydispersity may be in the range offrom about 1.3 to about 1.5.

In accordance with one preferred aspect of the invention, the reactionzone may comprise a loop reactor wherein the reaction admixture iscontinuously recirculated at a first volumetric flow rate, and saidfeedstock and said catalyst composition are continuously introduced at acombined second volumetric flow rate. The ratio of said first volumetricflow rate to said second volumetric flow rate may desirably range fromabout 20:1 to about 50:1, may preferably range from about 25:1 to about40:1 and ideally may range from about 28:1 to about 35:1. In order toachieve the benefits of the invention, the ratio of said firstvolumetric flow rate to said second volumetric flow rate may preferablybe such that the concentrations of ingredients in the reaction admixtureremain essentially constant and/or such that essentially isothermalconditions are established and maintained in said reaction admixture.

The feedstock and the catalyst composition may be premixed andintroduced into the reaction zone together as a single stream at saidsecond volumetric flow rate. Alternatively, the feedstock and thecatalyst composition may be introduced into the reaction zone separatelyas two respective streams, the flow rates of which together add up tosaid second volumetric flow rate.

To achieve the desired results of the invention, the reactorconfiguration, the properties of the reaction mixture, and the firstvolumetric flow rate may be such that turbulent flow is maintained insaid reaction zone. In particular, the system may be such that aReynolds number of at least about 2000 is achieved and maintained insaid reaction zone. The system may also be such that a heat transfercoefficient (U) of at least about 50 Btu/min ft²° F. is achieved andmaintained in said reaction zone. To this end, the reactor maypreferably be the tube side of a shell-and-tube heat exchanger.

In further accordance with the concepts and principles of the invention,the feed stock may generally comprise at least about 30% by weight ofisobutylene, with the remainder being non-reactive hydrocarbon diluents.

In a more specific sense, the invention may provide a liquid phasepolymerization process for preparing polyisobutylene having an averagemolecular weight in the range of from about 500 to about 5000 and avinylidene content of at least 70%. The process may comprise providingboth a feedstock comprising isobutylene and a separate catalystcomposition made up of a complex of BF₃ and a C₁ to C₈ primary alcohol.The molar ratio of BF₃ to alcohol in said complex may desirably be inthe range of from about 0.5:1 to about 2:1. The feedstock and thecatalyst composition may be introduced separately or together as asingle stream into a residual reaction mixture in a reaction zone, andthe residual reaction mixture, the feedstock and the catalystcomposition may be intimately intermixed to present an intimatelyintermixed reaction admixture in said reaction zone. The introduction ofthe catalyst complex into the reaction admixture may preferably becontrolled so that about 0.1 to about 10 millimoles of BF₃ areintroduced for each mole of isobutylene introduced with the feedstock.The intimately intermixed condition of the reaction admixture shouldpreferably be maintained and the temperature thereof kept at about 0° C.or above while the admixture is in the reaction zone, whereby theisobutylene in the admixture undergoes polymerization to form saidpolyisobutylene. Thereafter, a product stream comprising thepolyisobutylene product may be withdrawn from the reaction zone. Theintroduction of said feedstock into the reaction zone and the withdrawalof the product stream from the reaction zone may preferably be such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is no greater than about 4 minutes.

Even more desirably, the invention may provide a liquid phasepolymerization process for preparing polyisobutylene having an averagemolecular weight in the range of from about 950 to about 1050, apolydispersity within the range of from about 1.3 to about 1.5, and avinylidene content of at least about 80%. In accordance with thispreferred aspect of the invention, the process comprises providing botha feedstock made up of at least about 40% by weight isobutylene and aseparate catalyst composition made up of a complex of BF₃ and methanol,wherein the molar ratio of BF₃ to methanol in the complex ranges fromabout 0.5:1 to about 1:1. The feedstock and the catalyst composition areintroduced either separately or together into a residual reactionmixture in a reaction zone. The residual reaction mixture, the feedstockand the catalyst composition are intimately intermixed by turbulent flowwithin said reaction zone, whereby an intimately intermixed reactionadmixture is present in the reaction zone. Preferably, the catalystcomplex is introduced into the reaction admixture at a rate such thatabout 0.5 to about 2 millimoles of BF₃ are introduced for each mole ofisobutylene introduced in the feedstock. The intimately intermixedcondition of the reaction admixture is maintained and the temperaturethereof is kept at about 0° C. or more while the same is in saidreaction zone, whereby the isobutylene therein is caused to undergopolymerization to form said polyisobutylene. A product stream comprisingsaid polyisobutylene is withdrawn from said reaction zone. In accordancewith the invention, the introduction of feedstock into the reaction zoneand the withdrawal of product stream therefrom are controlled such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is within the range of from about 45 to about 90 seconds.

BRIEF DESCRIPTION OF THE DRAWING

FIG. 1 is a schematic illustration of a reactor in the form of amulti-pass shell and tube heat exchanger which is useful for carryingout the improved process of the invention; and

FIG. 2 is a schematic illustration of an alternative reactor in the formof a single pass shell and tube exchanger which is also useful forcarrying out the improved process of the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

In accordance with one very important embodiment of the presentinvention, an improved liquid phase process is provided for theefficient and economical production of PIB. In accordance with thisembodiment of the invention, an isobutylene containing feedstock streamis contacted in a reaction zone with a catalyst which facilitates thepolymerization reaction. Appropriate reaction conditions are provided inthe reaction zone. After an appropriate residence time, a PIB containingproduct stream is withdrawn from the reaction zone. As mentioned above,many techniques for conducting the reaction are known; however, from acommercial viewpoint it is always desirable to improve the efficiencyand economics of the process. With the foregoing in mind, the presentinvention provides an improved PIB producing process which may be easilycontrolled and manipulated to efficiently and economically provide arelatively low molecular weight, highly reactive PIB product.

The improved process of the present invention features the use of a BF₃catalyst which desirably may be complexed with a complexing agent whichappropriately alters the performance of the catalyst. Many otherpotentially useful catalysts are known to those of ordinary skill in therelated art field. In particular, many useful catalysts are described inthe prior patents referenced above. The complexing agent for thecatalyst, and in particular for the BF₃ catalyst, may be any compoundcontaining a lone pair of electrons, such as, for example, an alcohol,an ester or an amine. For purposes of the present invention, however,the complexing agent preferably may be an alcohol, desirably a primaryalcohol, preferably a C₁-C₈ primary alcohol and ideally methanol.

The molar ratio of BF₃ to complexing agent in the catalyst compositionmay generally be within the range of from approximately 0.5:1 toapproximately 5:1, desirably within the range of from approximately0.5:1 to approximately 2:1, and preferably within the range of fromapproximately 0.5:1 to approximately 1:1. Ideally, the catalystcomposition may simply be a 1:1 complex of BF₃ and methanol. In somepreferred embodiments of the invention, the molar ratio of BF₃ tocomplexing agent in said complex may be approximately 0.75:1.

The temperature in the reaction zone may generally and preferably begreater than 0° C., the reactor residence time may generally andpreferably be less than 4 minutes and the desired vinylidene (terminalunsaturation) content in the PIB product may preferably and generally begreater than about 70%. With these parameters, it is possible to operatethe process so as to achieve efficiencies and economies not previouslythought to be available. In accordance with the present invention, thecatalyst concentration and the BF₃/complexing ratio may be manipulatedas required to achieve the desired 70% vinylidene content with areaction temperature greater than 0° C. and a reactor residence time ofless than 4 minutes. Generally speaking, for PIB production the amountof the BF₃ catalyst introduced into the reaction zone should be withinthe range of from about 0.1 to about 10 millimoles for each mole ofisobutylene introduced into the reaction zone. Preferably, the BF₃catalyst may be introduced at a rate of about 0.5 to about 2 millimolesper mole of isobutylene introduced in the feedstock.

The process itself includes steps resulting in the intimate mixing ofthe isobutylene containing reactant stream and the catalyst complexand/or removal of heat during the reaction. The intimate mixing maydesirably be accomplished by turbulent flow. Turbulent flow alsoenhances heat removal. These conditions separately or together permitthe higher operating temperatures (e.g. >0° C.) and the shorter reactorresidence times (e.g. <4 minutes) provided by the invention. Theseimportant parameters may be achieved by causing the catalyzed reactionto take place in the tubes of a shell-and-tube heat exchanger at a flowrate which results in turbulent flow.

Many potentially valuable reactors are well known to the routineers inthe art to which the invention pertains. However, for purposes of onepreferred embodiment of the invention, the reactor may be a four-passshell-and-tube heat exchanger as shown in FIG. 1 where it is identifiedby the numeral 10. The reactor may, for example, have 80⅜-inch tubeswith a wall thickness of 0.022 inch, each thereby providing an internaltube diameter of 0.331 inch. The reactor may be three feet long and mayhave internal baffling and partitions to provide 4 passes with 20 tubesper pass. Such construction is well known in the heat exchanger andreactor arts and no further explanation is believed necessary.

In operation, in accordance with the preferred procedure for producinghighly reactive PIB, the isobutylene containing feedstock enters thereactor system through pipe 15 which is preferably located adjacent thebottom head 11 of reactor 10. Pipe 15 directs the feed stock into thesuction line 20 of a recirculation pump 25. The catalyst complex may beinjected into the reactor circulation system through pipe 30 locatedadjacent bottom head 11 of reactor 10. It should be noted here, that inaccordance with the principles and concepts of the invention, thecatalyst complex could just as well be injected separately into thereactor, in which case a separate catalyst pump might be required.

A catalyst modifier may be added to the feedstock via pipe 16 before thefeedstock enters the reactor system. The purpose of the modifier is toassist in controlling the vinylidene content of the PIB product. Thecatalyst modifier may be any compound containing a lone pair ofelectrons such as an alcohol, an ester or an amine. However, it ispointed out in this regard that if the amount of modifier is too great,the same may actually kill the catalyst. The feedstock containing themodifier enters the reactor system at the suction line 20 of thecirculation pump 25. The catalyst complex composition enters the reactorsystem via pipe 30 at a location downstream from pump 25 and adjacentthe first pass as shown in FIG. 1. The catalyst complex is preferably amethanol/BF₃ complex with a 1:1 molar ratio of BF₃ to methanol. Theamount of modifier added via line 16 may vary from 0 to about 1 mole foreach mole of BF₂ added as a complex via pipe 30.

Circulation pump 25 pushes the reaction mixture through line 35, controlvalve 40 and line 45 into the bottom head 11 of the reactor 10. A flowmeter 46 may be positioned in line 45 as shown. The reaction mixturetravels upwardly through pass 50, downwardly through pass 51, upwardlythrough pass 52 and downwardly through pass 53. As explained previously,each pass 50, 51, 52 and 53 may preferably include 20 separate tubes.For clarity, only a respective single tube is illustrated schematicallyin each pass in FIG. 1. These tubes are identified by the referencenumerals 50 a, 51 a, 52 a and 53 a. However, as discussed above, eachpass may preferably consist of a plurality, for example, 20 of theseindividual tubes.

It is to be noted here, that the reaction mixture should preferably becirculated through the tubes 50 a, 51 a, 52 a, 53 a of the reactor at aflow rate sufficient to obtain turbulent flow, whereby to achieveintimate intermixing between the catalyst complex and the reactants anda heat transfer coefficient appropriate to provide proper cooling. Inthis regard, the flow rate, the reaction mixture properties, thereaction conditions and the reactor configuration should be appropriateto produce a Reynolds number (Re) in the range of from about 2000 toabout 3000 and a heat transfer coefficient (U) in the range of fromabout 50 to about 150 Btu/min ft²° F. in the tubes of the reactor. Suchparameters may generally be obtained when the linear flow rate of atypical reaction mixture through a tube having an internal diameter of0.331 inch is within the range of from about 6 to 9 feet per second.

The circulating reaction mixture leaves reactor 10 via suction line 20.The recirculation rate in the system is preferably sufficiently high sothat the reactor, in essence, is a Continuous Stirred Tank Reactor(CSTR). In this same connection, and generally speaking, therecirculation rate of the reaction mixture should preferably be suchthat essentially steady state conditions are maintained in the reactor.It is pointed out in this latter regard that regardless of the systembeing in an unsteady or steady state, the design equations may bereduced to those of a CSTR when the recycle rate is sufficiently high.The reactor may also be of the type which is sometimes referred to as aloop reactor. With this system, which is only a preferred system sincethere are many other arrangements which would be apparent to those ofordinary skill in the art, the flow rate of the reactant mixture in thereactor may be adjusted and optimized independently of feed stockintroduction and product removal rates so as to achieve thoroughintermixing of the catalyst complex and the reactants and appropriatetemperature control.

A product exit line 55 may preferably be provided in top head 12 at apoint which is approximately adjacent the transition zone between thethird and fourth passes. Such positioning may be desirable to avoid anypotential for loss of unreacted isobutylene. Moreover, the positioningof the exit line 55 should be appropriate to facilitate bleeding of gasfrom the reactor during startup. A coolant may desirably be circulatedon the shell side of the reactor at a rate to remove heat of reactionand maintain the desired temperature in the reactor.

The product exiting the system via line 55 should be quenchedimmediately with a material capable of killing the catalyst, such as,for example, ammonium hydroxide. Thus, any potential rearrangement ofthe polymer molecule which would shift the double bond away from theterminal position is minimized. The high vinylidene isobutylene productmay then be directed to a work up system (not shown) where catalystsalts may be removed and the isobutylene product separated fromunreacted isobutylene and other undesirable contaminants such asdiluents, etc. These latter materials may then be recycled or divertedfor other uses employing known methodology.

With the described recirculation system, the rate of feedstockintroduction into the reaction mixture and the rate of product removalare each independent of the circulation rate. As will be appreciated bythose of ordinary skill in the art, the number of passes through thereactor and the size and configuration of the latter are simply mattersof choice. The feedstock and product withdrawal flow rates maypreferably be chosen such that the residence time of the reactionmixture within the reactor is 4 minutes or less, desirably 3 minutes orless, preferably 2 minutes or less, even more preferably 1 minute orless, and ideally less than 1 minute. From a commercial operatingviewpoint, the flow rate should be such that the residence time of thereaction mixture in the reactor is within the range of from about 45 toabout 90 seconds. In connection with the foregoing, the residence timeis defined as the total reactor system volume divided by the volumetricflow rate.

The recirculation flow rate, that is the flow rate of the reactionmixture in the system induced by the recirculation pump 25, iscontrolled, as described above, to achieve appropriate turbulence and/orheat transfer characteristics. This recirculation flow rate is often afunction of the system itself and other desired process conditions. Forthe system described above, the ratio of the recirculation flow rate tothe incoming feedstock flow rate (recycle ratio) should generally bemaintained in the range of from about 20:1 to about 50:1, desirably inthe range of from about 25:1 to about 40:1, and ideally in the range offrom about 28:1 to about 35:1. In particular, in addition to causingturbulence and providing an appropriate heat transfer coefficient, therecirculation flow rate of the reaction mixture should be sufficient tokeep the concentrations of the ingredients therein essentially constantand/or to minimize temperature gradients within the circulating reactionmixture whereby essentially isothermal conditions are established andmaintained in the reactor.

As mentioned above, the recycle ratios generally may be in the range offrom about 20:1 to about 50:1 when the desired product is highlyreactive PIB. Higher recycle ratios increase the degree of mixing andthe reactor approaches isothermal operation leading to narrower polymerdistributions. Lower recycle ratios decrease the amount of mixing in thereactor, and as a result, there is a greater discrepancy in thetemperature profiles. As the recycle ratio approaches zero, the designequations for the reactor reduce to those for a plug flow reactor model.On the other hand, as the recycle ratio approaches infinity, themodeling equations reduce to those for a CSTR. When CSTR conditions areachieved, both temperature and composition remain constant and thecomposition of the product stream leaving the reactor is identical tothe composition of the reaction mixture recirculating in the reactor.

Needless to say, after steady state or near steady state operation hasbeen established in the reactor, as the feedstock enters the system, anequal volume of product is pushed out of the reactor loop. Under CSTRconditions, the point at which the product stream is withdrawn isindependent of reactor geometry. However, the top of the third pass waschosen for this particular embodiment of the invention so any air ornon-condensable species in the reactor at start-up may conveniently bepurged. Also, it is preferred that the withdrawal point be as far aspossible from the point where fresh feedstock is introduced into thesystem just to make sure that conditions within the reactor haveachieved steady state operation and are therefore as stable as possible.

When highly reactive PIB is the desired product, the feedstock enteringthe system through line 15 may be any isobutylene containing stream suchas, but not limited to, isobutylene concentrate, dehydro effluent, or atypical raff-1 stream. These materials are described respectively belowin Tables 1, 2 and 3.

TABLE 1 Isobutylene Concentrate Ingredient Weight % C₃ component 0.00I-butane 6.41 n-butane 1.68 1-butene 1.30 I-butene 89.19 trans-2-butene0.83 cis-2-butene 0.38 1,3-butadiene 0.21

TABLE 2 Dehydro Effluent Ingredient Weight % C₃ components 0.38 I-butane43.07 n-butane 1.29 1-butene 0.81 I-butene 52.58 trans-2-butene 0.98cis-2-butene 0.69 1,3-butadiene 0.20

TABLE 3 Raff-1 Ingredient Weight % C₃ components 0.57 I-butane 4.42n-butane 16.15 1-butene 37.22 I-butene 30.01 trans-2-butene 8.38cis-2-butene 2.27 1,3-butadiene 0.37 MTBE 0.61

For commercial and process economies, the isobutylene content of thefeedstock generally should be at least about 30 weight %, with theremainder comprising one or more non-reactive hydrocarbon, preferablyalkane, diluents.

The desired product is a relatively low molecular weight, highlyreactive polyisobutylene. Thus, the polyisobutylene leaving the reactorby way of line 55 should have an average molecular weight which is lessthan about 10,000. Generally speaking, the produced isobutylene shouldhave an average molecular weight within the range of from about 500 toabout 5000, desirably from about 600 to about 4000, preferably fromabout 700 to about 3000, even more preferably from about 800 to about2000, and ideally from about 950 to about 1050. By carefully controllingthe various parameters of the process, it might even be possible toproduce a product wherein the average molecular weight is consistent atsome desired number, for example, 1000.

The polydispersity of the PIB may also be important. The termpolydispersity refers to the molecular weight distribution in a givenpolymer product and generally is defined as the ratio of the molecularweight of the highest molecular weight molecule to the molecular weightof the lowest molecular weight molecule. Polydispersity may becontrolled by carefully maintaining constant monomer concentrations andisothermal conditions within the reaction mixture. Generally speaking,it is desirable that the polydispersity be as low as possible in orderto diminish the content of unwanted relatively low or high molecularweight polyisobutylenes in the product and thus improve the quality ofthe latter. By following the concepts and principles of the presentinvention, it has been found that the polydispersity of the product maybe controlled at no more than about 2.0. Preferably, through the use ofthe invention, a polydispersity of no more than about 1.65 may beachieved. Even more desirably, the polydispersity may be controlled soas to be within the range of from about 1.3 to about 1.5.

The polyisobutylene product obtained through the use of the presentinvention should generally have a terminal (vinylidene) unsaturationcontent of at least about 70%. That is to say, at least about 70% of thedouble bonds remaining in the polymerized product should preferably bein a terminal position. Ideally the vinylidene content should be no lessthan about 80% or even higher. However, vinylidene content is indirectlyrelated to conversion rates. That is to say, the higher the conversionrate, the lower the vinylidene content. Moreover, vinylidene content isdirectly related in the same way to molecular weight. Accordingly, ineach process a balance is required between molecular weight, conversionrate and vinylidene content.

EXAMPLE 1

Using the principles and concepts of the invention, a reactor such asthe reactor illustrated in FIG. 1, was used to produce a low molecularweight, highly reactive polyisobutylene. The feedstock was essentiallythe same as shown above in Table 1, and the coolant circulated on theshell side of the reactor was a mixture of 35 weight % methanol and 65weight % water. The inlet coolant temperature was 32° F. A 1:1BF₃/methanol complex catalyst was used. All pertinent reactor data anddimensions are set forth below in Table 4.

TABLE 4 Feedstock flow rate 1.7 gpm Recirculation flow rate 50 gpmFeedstock density 5 lb/gal Conversion 63 wt % Concentration ofisobutylene in 92 wt % feedstock ΔH_(reaction) 398 Btu/lb μ reactionmixture 4.5 cP = 0.0030 lb/ft sec Cp of reaction mixture 0.46 Btu/lb °F. Reaction effective density 44.9 lb/ft³ Thermal conductivity 0.075Btu/hr ft ° F. Total volume of reactor recirculation 390.2 in³ Residencetime 59.6 seconds Linear velocity inside tubes 9.3 ft/sec Reynoldsnumber 3180 Surface area of tubes 23.6 ft² Heat generated 1961 Btu/minΔT_(lm) 37.3° F. Heat flux 83.2 Btu/min ft² U 133.7 Btu/min ft² ° F. Cpof coolant 0.86 Btu/lb ° F. Density of coolant 7.70 lb/gal Coolant flowrate 39.3 gpm ΔT coolant 8.0° F. Heat removed 2074 Btu/min

The composition of the product thus obtained is as set forth below inTable 5.

TABLE 5 Crude Polyisobutylene Product Ingredient Weight % C₃ components0.00 I-butane 6.41 n-butane 1.68 1-butene 1.30 I-butene 33.00trans-2-butene 0.83 cis-2-butene 0.38 1,3-butadiene 0.21 polyisobutylene56.19

Again it is to be noted that one of the main objectives in accordancewith the invention is to provide a flow rate through the reactor andother parameters such that the reaction mixture is in a generallyconstant state of turbulent flow during the reaction. Turbulent flowresults in a twofold augmentation of the overall process. First,turbulent flow results in intimate intermixing of the contents of thereactor to enhance the kinetics of the reaction. Second, turbulent flowresults in an enhancement of the tube side heat transfer coefficient tothereby improve the removal of the heat of the reaction. These resultsmay be achieved by conducting the reaction on the tube side of ashell-and-tube heat exchanger reactor and circulating a coolant on theshell side.

The foregoing description concerns methodology which permits the PIBpolymerization reaction to be conducted at higher temperatures and atlower residence times than current processes. In accordance with thisembodiment of the present invention, a stable BF₃ catalyst system(BF₃/methanol) may be used. Moreover, an improved turbulent loop reactorconfiguration including a heat exchanger to effect simultaneous heatremoval is advantageously employed. The turbulent flow also enablesintimate mixing of the two-phase reaction system.

In addition to highly reactive PIB, the process of the inventionprovides an improved process for preparing oligomers and highermolecular weight polymers from olefinic precursors. In general, theprocess of the invention may be used to produce conventional PIB, lowmolecular weight oligomers of branched olefins, preferably isobutylene,oligomers and higher molecular weight polymers of linear C₃-C₁₅ alphaolefins, and oligomers and higher molecular weight polymers of C₄-C₁₅reactive non-alpha olefins. In accordance with this aspect of theinvention, and particularly where the desired product is a relativelylow molecular weight (<350 and perhaps even <250) oligomer, the catalystcomplex is desirably stable, even under the relatively higher reactiontemperatures needed for the production of oligomeric olefinic products.

Examples of processes for production of relatively low molecular weightoligomers of olefinic monomeric components are set forth below. In theseexamples, a loop reactor as illustrated in FIG. 2 is utilizedadvantageously. As illustrated in FIG. 2, the reactor 100 may consist ofa single tube 102 surrounded by a heat exchanger shell 104. In all otheressential aspects, the recirculation system may preferably be the sameas described in connection with the reactor 10 of FIG. 1, except that arecirculation line 106 is provided to return the recirculating residualmixture from the top of reactor tube 102 to the pump suction line 20.The exit line 55 is connected directly to recirculation line 106 asshown.

EXAMPLE 2

A stream containing 2.19 wt % Isobutane, 61.5 wt % n-butane, 0.64 wt %1-butene, 28.18 wt % trans-2-butene and 7.49 wt % cis-2-butene (35.66 wt% 2-butene) is introduced into the a loop reactor system of FIG. 2 viafeed line 15 at a rate of 156 ml/min (93.6 g/min). A catalyst complexcontaining BF3/methanol complex (one mole of BF3 to one mole ofmethanol) is fed to the reactor at a rate of 8 ml/min (10.4 g/min). Thereaction temperature is maintained constant at 90° F. All pertinentreactor data and dimensions are set forth below in Table 6. The reactoreffluent exits the top of the reaction loop via line 55 and is fed intoa decant (not shown) were the catalyst is preferably separated out fromthe organic layer. A portion of the catalyst may then be recycled backto the reactor lowering the amount of fresh catalyst required. Theproduct coming out of the decant overhead is mixed with NH₄OH to quenchany remaining catalyst in the organic phase and is sent to a seconddecant. The products is washed twice more with water and decanted toremove the last traces of catalyst. The oligomer product composition isgiven in Table 7.

TABLE 6 HC flow rate 0.0412 gpm Pump around flow rate 1.5 gpm HC density5 lb/gal % Conversion 51 wt % % 2-butene in feed stock 36.55 wt %ΔH_(rxn) 318 Btu/lb μ 0.6 cP = 0.0004/ lb/ft-s Cp 0.46 Btu/lb-° F.Reactor OD 0.375 in Reactor wall thickness 0.035 in Reactor ID 0.305 inReactor length 10.5 ft Reactor volume 9.2 in³ # of tubes 1 # of passes 1Residence time 58.02 seconds Linear velocity 6.59 ft/s Surface area 1.03ft² Heat generated 12.2 Btu/min ΔT_(lm) 3.0° F. Heat flux 11.8Btu/min-ft² U 237.0 Btu/hr-ft²-° F. Re 15531

TABLE 7 C₈  7.9 wt % C₁₂ 29.8 wt % C₁₆ 35.9 wt % C₂₀ 16.1 wt % C₂₀₊ 10.3wt %

EXAMPLE 3

A stream containing 94.0 wt % 1-decene and 6.0 wt % C₁₀-isomers was fedinto the loop reactor of FIG. 2 at a rate of 10 ml/min (7.4 g/min). Acatalyst complex containing BF3/methanol complex (one mole of BF3 to onemole of methanol) was fed to the reactor at a rate of 1 ml/min (1.3g/min). The reaction was held at a constant temperature of 70° F. Allpertinent reactor data and dimensions are given in Table 8. Both thereactor setup and downstream catalyst removal steps are identical toExample 2. The product stream contained about 59.8 wt % of C₂₀ oligomersand about 40.2 wt % of C₃₀ oligomers.

TABLE 8 HC flow rate 0.00264 gpm Pump around flow rate 1.5 gpm HCdensity 6.2 lb/gal % Conversion 90 wt % % 1-decene in feed stock 94 wt %ΔH_(rxn) 318 Btu/lb μ 1.2 cP = 0.0008 lb/ft s Cp 0.50 Btu/lb° F. ReactorOD 0.375 in Reactor wall thickness 0.035 in Reactor ID 0.305 in Reactorlength 10.5 ft Reactor volume 9.2 in³ # of tubes 1 # of passes 1Residence time 905.13 seconds Linear velocity 6.59 ft/s Surface area1.03 ft² Heat generated 4.4 Btu/min ΔT_(lm) 1.2° F. Heat flux 4.3Btu/min-ft² U 213.2 Btu/hr-ft²-° F. Re 9604.4

As can be seen from the foregoing examples, the invention provides aprocess for preparing a polyolefin product having preselectedproperties. In accordance with the invention, the process advantageouslyemploys a stable complex of BF₃ and a complexing agent therefor. Theresidual reaction mixture in the reaction zone is recirculated at arecirculation rate sufficient to cause intimate intermixing of thereaction mixture. The rate is also such that the heat of reaction isremoved from the reaction mixture at a rate calculated to provide asubstantially constant reaction temperature in the reaction mixturewhile the same is recirculating in the reaction zone. The introductionof the feedstock and the withdrawal of the product stream arecontrolling such that the residence time of the olefinic componentsundergoing polymerization in the reaction zone is appropriate forproduction of the desired polyolefin product.

Although the foregoing text and examples have focused on processeswherein a single monomer is included in the feedstock, it will beapparent to the routineers in the olefin polymerization art that inaccordance with the principles and concepts of the present invention,the feedstock may desirably, at times, include two or more monomers soas to produce useful copolymeric products.

1. A single-stage liquid phase polymerization process for preparing apolyolefin product having preselected properties, said processcomprising: providing a liquid feedstock comprising at least oneolefinic component; providing a catalyst composition comprising acomplex of BF₃ and a complexing agent therefor which comprises analiphatic alcohol having a primary hydroxyl group and no hydrogen on abeta carbon, said complex thereby being stable at temperatures greaterthan 0° C.; introducing said feedstock and said catalyst compositioninto a residual reaction mixture in a loop reactor reaction zone;recirculating the residual reaction mixture with the feedstock and thecatalyst composition therein in said zone at a recirculation ratesufficient to cause intimate intermixing of the residual reactionmixture, the feedstock and the catalyst composition to thereby present arecirculating, intimately intermixed reaction admixture of the residualreaction mixture, the feedstock and the catalyst composition in saidreaction zone; maintaining the recirculating intimately intermixedreaction admixture in its intimately intermixed condition and removingheat of reaction from the reaction admixture at a rate calculated toprovide a substantially constant reaction temperature greater than 0° C.in the reaction admixture while the reaction admixture is recirculatingin said reaction zone; withdrawing a product stream comprisingpolyolefin product and catalyst composition from said reaction zone; andcontrolling the introduction of said feedstock into said reaction zoneand the withdrawal of said product stream from the reaction zone suchthat the residence time of the olefinic components undergoingpolymerization in the reaction zone is appropriate for production ofsaid polyolefin product.
 2. A process as set forth in claim 1, whereinsaid reaction zone comprises a tube side of a shell-and-tube heatexchanger.
 3. A process as set forth in claim 1, further comprisingadding a quenching material capable of killing the catalyst to thewithdrawn product stream.
 4. A process as set forth in claim 1, whereinsaid residence time is no greater than 4 minutes.
 5. A process as setforth in claim 1, wherein said complexing agent comprises a C₁-C₈aliphatic alcohol.
 6. A process as set forth in claim 5, where saidalcohol comprises methanol.
 7. A process as set forth in claim 1,wherein the molar ratio of BF₃ to complexing agent in said complexranges from approximately 0.5:1 to approximately 1:1.
 8. A process asset forth in claim 1, wherein said feedstock and said catalystcomposition are premixed and introduced into the reaction zone togetheras a single stream.
 9. A process as set forth in claim 1, wherein saidfeedstock and said catalyst composition are introduced into the reactionzone separately as two streams.
 10. A process as set forth in claim 1,wherein said residual reaction mixture is recirculated at a rate suchthat a Reynolds number of at least 2000 is maintained in said reactionzone.
 11. A process as set forth in claim 1, wherein a heat transfercoefficient U of at least 50 Btu/min ft²° F. is maintained in saidreaction zone.
 12. A process as set forth in claim 1, wherein said feedstock comprises at least 30% by weight of said olefinic component.
 13. Aprocess as set forth in claim 12, wherein said olefinic componentcomprises isobutylene.
 14. A process as set forth in claim 1, whereinthe olefinic component comprises isobutylene and the polyolefin productcomprises polyisobutylene.
 15. A process as set forth in claim 14,wherein said polyisobutylene has a vinylidene content of at least 50%.16. A process as set forth in claim 1, wherein the olefinic componentcomprises a C₃ to C₁₅ linear alpha olefin.
 17. A process as set forth inclaim 1, wherein the olefinic component comprises a C₄ to C₁₅ non-alphaolefin.
 18. A process as set forth in claim 2, further comprising addinga quenching material capable of killing the catalyst to the withdrawnproduct stream.
 19. A process as set forth in claim 2, wherein saidresidual reaction mixture is recirculated at a rate such that a Reynoldsnumber of at least 2000 is maintained in said reaction zone.
 20. Aprocess as set forth in claim 3, wherein said residual reaction mixtureis recirculated at a rate such that a Reynolds number of at least 2000is maintained in said reaction zone.
 21. A single-stage liquid phasepolymerization process for preparing a polyolefin product havingpreselected properties, said process comprising: providing a liquidfeedstock comprising at least one olefinic component; providing acatalyst composition comprising a complex of BF₃ and a complexing agenttherefor, said complexing agent comprising a glycol, said complex beingstable at temperatures needed to produce said polyolefin product;introducing said feedstock and said catalyst composition into a residualreaction mixture in a loop reactor reaction zone; recirculating theresidual reaction mixture with the feedstock and the catalystcomposition therein in said zone at a recirculation rate sufficient tocause intimate intermixing of the residual reaction mixture, thefeedstock and the catalyst composition to thereby present arecirculating, intimately intermixed reaction admixture of the residualreaction mixture, the feedstock and the catalyst composition in saidreaction zone; maintaining the recirculating intimately intermixedreaction admixture in its intimately intermixed condition and removingheat of reaction from the reaction admixture at a rate calculated toprovide a substantially constant reaction temperature in the reactionadmixture while the reaction admixture is recirculating in said reactionzone, said constant reaction temperature being at a level appropriatefor causing olefinic components introduced in said feedstock to undergopolymerization to form said polyolefin product in the presence of saidcatalyst composition; withdrawing a product stream comprising polyolefinproduct and catalyst composition from said reaction zone; andcontrolling the introduction of said feedstock into said reaction zoneand the withdrawal of said product stream from the reaction zone suchthat the residence time of the olefinic components undergoingpolymerization in the reaction zone is appropriate for production ofsaid polyolefin product.
 22. A process as set forth in claim 21, wheresaid complexing agent comprises ethylene glycol.